Method of producing diamines and polyamines of the diphenylmethane series at different production capacities

ABSTRACT

The invention relates to a method for producing diamines and polyamines of the diphenylmethane series, by condensing aniline and formaldehyde followed by an acid-catalysed rearrangement at different production capacities. Adapting the respective molar ratios of the total used aniline to the total used formaldehyde and of the total used acid catalyst to the total used aniline allows the rearrangement reaction to be fully completed despite the change in dwell time inevitably associated with a change in production capacity and ensures that no undesired fundamental changes occur in the product composition.

The present invention relates to a process for preparing di- and polyamines of the diphenylmethane series by condensation of aniline and formaldehyde followed by acid-catalyzed rearrangement at different production capacities. Adjusting the respective molar ratios of total aniline used to total formaldehyde used and of total acidic catalyst used to total aniline used achieves running of the rearrangement reaction to completion in spite of the change in residence time inevitably associated with the altered production capacity and avoiding unwanted significant changes in the product composition.

The preparation of di- and polyamines of the diphenylmethane series (MDA) by reaction of aniline with formaldehyde in the presence of acidic catalysts is known in principle. In the context of the present invention, di- and polyamines of the diphenylmethane series are understood to mean amines and mixtures of amines of the following type:

n here is a natural number ≥2. The compounds of this type in which n=2 are referred to hereinafter as diamines of the diphenylmethane series or diaminodiphenylmethanes (MMDA hereinafter; “monomer MDA”). Compounds of this type in which n>2 are referred to in the context of this invention as polyamines of the diphenylmethane series or polyphenylene polymethylene polyamines (PMDA hereinafter; “polymer MDA”). Mixtures of the two types are referred to as di- and polyamines of the diphenylmethane series (for the sake of simplicity referred to hereinafter as MDA). The corresponding isocyanates that can be derived in a formal sense by replacing all NH₂ groups with NCO groups from the compounds of the formula (I) are accordingly referred to as diisocyanates of the diphenylmethane series (MMDI hereinafter), polyisocyanates of the diphenylmethane series or polyphenylene polymethylene polyisocyanates (PMDI hereinafter) or di- and polyisocyanates of the diphenylmethane series (MDI hereinafter). The higher homologs (n>2), both in the case of the amine and in the case of the isocyanate, are generally always present in a mixture with the dimers (n=2), and so only two product types are of relevance in practice, the pure dimers (MMDA/MMDI) and the mixture of dimers and higher homologs (MDA/MDI). The position of the amino groups on the two phenylene groups in the dimers (para-para; ortho-para and ortho-ortho) is specified hereinafter only when it is important. For the sake of simplicity, this is done in the X,Y′-MDA form (4,4′-, 2,4′- or 2,2′-MDA), as is customary in the literature. The same is true of MDI (identification of the isomers as X,Y′-MDI (4,4′-, 2,4′- or 2,2′-MDI).

Industrially, the di- and polyamine mixtures are converted predominantly by phosgenation to the corresponding di- and polyisocyanates of the diphenylmethane series.

The continuous or partly discontinuous preparation of MDA is disclosed, for example, in EP-A-1 616 890, U.S. Pat. No. 5,286,760, EP-A-0 451 442 and WO-A-99/40059. The acidic condensation of aromatic amines and formaldehyde to give di- and polyamines of the diphenylmethane series proceeds in multiple reaction steps. In what is called the “aminal process”, in the absence of an acidic catalyst, formaldehyde is first condensed with aniline to give what is called aminal, with elimination of water. This is followed by the acid-catalyzed rearrangement to MDA in a first step to give para- or ortho-aminobenzylaniline. The aminobenzylanilines are converted to MDA in a second step. Main products of the acid-catalyzed reaction of aniline and formaldehyde are the diamine 4,4′-MDA, its positional isomers 2,4′-MDA and 2,2′-MDA and the higher homologs (PMDA) of the various diamines. In what is called the “neutralization process”, aniline and formaldehyde are converted in the presence of an acidic catalyst directly to aminobenzylanilines, which are then rearranged further to give the bicyclic MMDA isomers and higher polycyclic PMDA homologs. The present invention relates to the aminal process.

Irrespective of the process variant for preparation of the acidic reaction mixture, the workup thereof, according to the prior art, is initiated by neutralization with a base. This neutralization is typically effected at temperatures of, for example, 90° C. to 100° C. without addition of further substances (cf. H. J. Twitchett, Chem. Soc. Rev. 3 (2), 223 (1974)). It can alternatively be effected at a different temperature level in order, for example, to accelerate the degradation of troublesome by-products. Hydroxides of the alkali metal and alkaline earth metal elements are suitable as bases. Preference is given to using sodium hydroxide solution.

After the neutralization, the organic phase is separated from the aqueous phase in a separating vessel. The organic phase which comprises crude MDA and remains after removal of the aqueous phase is subjected to further workup steps, for example washing with water (base washing) in order to wash residual salts out of the crude MDA. Finally, the crude MDA thus purified is freed of excess aniline, water and other substances present in the mixture (e.g. solvents) by suitable methods, for example distillation, extraction or crystallization. The workup which is customary according to the prior art is disclosed, for example, in EP-A-1 652 835, page 3 line 58 to page 4 line 13, or EP-A-2 103 595, page 5 lines 21 to 37.

International patent application WO 2014/173856 A1 provides a process for preparing di- and polyamines of the diphenylmethane series by converting aniline and formaldehyde in the absence of an acidic catalyst to aminal and water, removing the aqueous phase and processing the organic aminal phase further to give the di- and polyamines of the diphenylmethane series, in which use of a coalescence aid in the phase separation of the process product obtained in the aminal reaction reduces the proportion of water and hence also of water-soluble impurities in the organic phase comprising the aminal. The di- and polyamines of the diphenylmethane series that are obtained by acid-catalyzed rearrangement and workup after further processing of the aminal phase are of excellent suitability for preparation of the corresponding isocyanates.

The quality of a reaction process for preparation of MDA is defined firstly by the content of unwanted secondary components and impurities in the crude product that can arise from improper conduct of the reaction. Secondly, the quality of a reaction process is defined in that the entire process can be operated without technical production outage or problems that necessitate intervention in the process, and that losses of feedstocks are prevented or at least minimized

Although the prior art processes described succeed in preparing MDA with a high yield and without loss of quality in the end products, the only processes described are in the normal state of operation. Only a few publications are concerned with states outside normal operation:

International patent application WO 2015/197527 A1 relates to a process for preparing di- and polyamines of the diphenylmethane series (MDA), to a plant for preparation of MDA and to a method of operating a plant for preparation of MDA. The invention enables optimization of production shutdowns in the operation of the MDA process with regard to time taken and optionally also with regard to energy and material consumption by means of what is called a circulation mode in individual plant components. During an interruption in the process or interruption of the operation of individual plant components, there is no introduction of formaldehyde into the reaction, and the plant components that are not affected by an inspection, repair or cleaning measure are operated in what is called circulation mode. What this achieves, among other effects, is that only the plant component in question must be shut down for the period of the measure, which may be advantageous in terms of productivity and economic viability of the process and the quality of the products prepared.

International patent application WO 2015/197519 A1 relates to a process for preparing diamines and polyamines of the diphenylmethane series, in which care is taken during the running-down of the production process that an excess of aniline over formalin is ensured.

International patent application WO 2015/197520 A1 relates to a process for preparing diamines and polyamines of the diphenylmethane series (MDA) from aniline and formaldehyde, in which care is taken during the start-up procedure to ensure that there is a sufficient excess of aniline over formaldehyde which is at least 105% of the molar ratio of aniline to formaldehyde wanted for the target formulation of the MDA to be produced.

Changes in the target production capacity (also called “change in load”) during continuous production (i.e. with a starting state and an end state, in which MDA is produced) are not taken into account here either. Since aniline is typically used in stoichiometric excess, the production capacity of a given plant for preparation of MDA is defined by the formaldehyde feed which, in the context of this invention, is referred to as (formaldehyde) load.

Changes in load are recurrent plant states that have a considerable effect on the economic (and environmentally benign in terms of energy consumption) operation of a continuously operating plant. Since, for a given continuously operated production plant, a change in load is inevitably associated with a change in the residence time of the reaction mixture in the reaction space available, as well as purely economic, environmental and operational challenges, there may additionally be an unwanted change in the composition of the product. This is because the exact composition of the MDA obtained (especially the isomer ratio of MMDA and the ratio of MMDA to PMDA) is also crucially dependent on the residence time of the reaction mixture, and so, in the event of improper procedure, a significantly different product can be obtained after a change in load, even though this was not intended.

It would thus be desirable to have available a process for preparing di- and polyamines of the diphenylmethane series in which it is possible by simple measures to configure changes in load in the operation of the MDA process such that they run in an optimized manner (for instance with regard to yield, time taken, energy consumption and avoidance of process-related problems such as caking or blockages in apparatus) in economic, environmental and operational aspects with avoidance of unwanted changes in the product composition.

According to the invention, this can be accomplished by a process according to patent claim 1. This process, as will be elucidated in detail hereinafter, is more particularly characterized in that in the event of an intended increase in load the molar ratio of total aniline used to total formaldehyde used, n(1)/n(2) (in the literature also referred to as A/F ratio), is lowered and the molar ratio of total acidic catalyst used to total aniline used, n(7)/n(1) (in the literature also as protonation level=n(7)/n(1)·100%), is increased and in that in the event of a decrease in load the reverse procedure is followed.

All configurations of the present invention relate to a process for preparing di- and polyamines of the diphenylmethane series from aniline (1) and formaldehyde (2) in a production plant (10 000), where the molar ratio of total aniline used (1) to total formaldehyde used (2), n(1)/n(2), is always greater than 1.6, comprising the steps of:

(A-I) reacting aniline (1) and formaldehyde (2) in the absence of an acidic catalyst in a reactor (1000, aminal reactor) to obtain a reaction mixture (4) comprising an aminal (3), and then at least partly separating an aqueous phase (6) from the reaction mixture (4) in a removal unit (2000, also called aminal separator) to obtain an organic phase (5) comprising the aminal (3);

(A-II) contacting the organic phase (5) which comprises the aminal and is obtained in step (A-I) with an acidic catalyst (7) in a reactor cascade (3000) composed of i reactors j connected in series (3000-j=1, 3000-j=2, . . . , 3000-j=i; hereinafter 3000-1, 3000-2, . . . , 3000-i for short), where i is a natural number from 2 to 10 (acid-catalyzed rearrangement), wherein

-   -   the first reactor (3000-1) in flow direction is operated at a         temperature T₃₀₀₀₋₁ in the range from 25.0° C. to 65.0° C.,         preferably in the range from 30.0° C. to 60.0° C., and is         charged with stream (5) and acidic catalyst (7) and optionally         with further aniline (1) and/or further formaldehyde (2),     -   every reactor downstream in flow direction (3000-2, . . . ,         3000-i) is operated at a temperature of more than 2.0° C. above         T₃₀₀₀₋₁ and is charged with the reaction mixture obtained in the         reactor immediately upstream;

(B) isolating the di- and polyamines of the diphenylmethane series from the reaction mixture (8-i) obtained from step (A-II) in the last reactor (3000-i) by at least

(B-I) adding a stoichiometric excess of base (9), based on the total amount of acidic catalyst used (7), to the reaction mixture (8-i) obtained in the last reactor (3000-i) in step (A-II) in a neutralization unit (4000, preferably stirred neutralization vessel) to obtain a reaction mixture (10), especially observing a molar ratio of base (9) to total acidic catalyst used (7) in the range from 1.01 to 1.30 (neutralization);

(B-II) separating the reaction mixture (10) obtained in step (B-I) in a separation unit (5000, neutralization separator) into an organic phase (11) comprising di- and polyamines of the diphenylmethane series and an aqueous phase (12).

The expressions “total aniline used” and “total formaldehyde used” each refer to the total amount of these reactants used in the process according to the invention, i.e. including any proportions additionally added after step (A-I).

The organic phase (11) can, as described in detail hereinafter, be subjected to further workup.

According to the invention, the abovementioned steps are conducted continuously, meaning that the respective reactants are supplied continuously to the apparatus assigned to the respective step and the products are withdrawn continuously therefrom.

According to the invention—in simplified terms; details are elucidated further down—the procedure is thus that, in the event of a change in load, the temperature in the reactors 3000-2, . . . , 3000-i is kept essentially constant and the molar ratio of total acidic catalyst used to total aniline used (n(7)/n(1) hereinafter) and the molar ratio of total aniline used to total formaldehyde used (n(1)/n(2) hereinafter) is adjusted such that, in spite of altered residence time, the rearrangement reaction runs to completion and unwanted changes in the product composition and increased by-product formation as a consequence of the change in load are largely avoided. According to the invention, the temperature in each reactor j of the reactor cascade (3000) after the change in load is essentially the same as before the change in load, although, in the case of the first reactor (3000-1) in flow direction, slightly greater deviations from the temperature before the change in load can be permitted during the change in load. These slightly greater tolerances in the temperature during the change in load allow easier control over the change in load in terms of control technology. The tolerances permitted in accordance with the invention, which are still to be elucidated in detail further down, however, are still within such a scope that operational problems are largely to completely avoided.

This procedure allows advantageous configuration of changes in load in economic, environmental (from the point of view of energy consumption) and operational aspects.

Since the invention relates to changes in load, i.e. to different states of operation of a production plant, the following definitions of terms that are used hereinafter are helpful:

Starting state A of the production plant with a given production capacity defined by the amount of formaldehyde introduced:

-   -   mass flow rate m₁ of total aniline used in the starting state:         m₁(A)≠0 [e.g. in t/h],     -   mass flow rate m₂ of total formaldehyde used in the starting         state: m₂(A)≠0 [e.g. in t/h],     -   mass flow rate m₇ of total acidic catalyst used in the starting         state: m₇(A)≠0 [e.g. in t/h],     -   molar ratio n(1)/n(2) of total aniline used (1) to total         formaldehyde used (2) in the starting state: n(1)/n(2)(A) and     -   molar ratio n(7)/n(1) of total acidic catalyst used to total         aniline used in the starting state: n(7)/n(1)(A).

End state E of the production plant with a target production capacity defined by the amount of formaldehyde introduced (with introduction of less or more formaldehyde compared to the starting state):

-   -   mass flow rate m₁ of total aniline used in the end state:         m₁(E)≠0 [e.g. in t/h],     -   mass flow rate m₂ of total formaldehyde used in the end state:         m₂(E)≠0 [e.g. in t/h],     -   mass flow rate m₇ of total acidic catalyst used in the end         state: m₇(E)≠0 [e.g. in t/h],     -   molar ratio n(1)/n(2) of total aniline used (1) to total         formaldehyde used (2) in the end state: n(1)/n(2)(E) and     -   molar ratio n(7)/n(1) of total acidic catalyst used to total         aniline used in the end state: n(7)/n(1)(E).

Transition state T between starting state and end state (corresponding to the period of the change in load):

-   -   mass flow rate m₁ of total aniline used in the transition state:         m₁(T)≠0 [e.g. in t/h],     -   mass flow rate m₂ of total formaldehyde used in the transition         state: m₂(T)≠0 [e.g. in t/h],     -   mass flow rate m₇ of total acidic catalyst used in the end         state: m₇(E)≠0 [e.g. in t/h],     -   molar ratio n(1)/n(2) of total aniline used (1) to total         formaldehyde used (2) in the transition state: n(1)/n(2)(T),     -   molar ratio n(7)/n(1) of total acidic catalyst used (7) to total         aniline used in the transition state: n(7)/n(1)(T).

According to the invention, the transition state begins as soon as m₂ is increased or lowered proceeding from m₂(A) (this time is referred to hereinafter as t₁), and it ends as soon as m₂ reaches the target value for m₂(E) (this time is referred to hereinafter as t₂).

On the basis of the minimum value for n(1)/n(2) of 1.6 used in the process of the invention, the yield of MDA target product is limited by the flow rate of total formaldehyde used (2) which is introduced. The maximum possible flow rate of total formaldehyde used (2) for preparation of a particular product type (i.e. a mixture of di- and polyamines of the diphenylmethane series characterized by a particular isomer composition of MMDA and a particular ratio of MMDA to PMDA) under the given boundary conditions of the production plant (10 000):

-   -   m₂(N) [e.g. in t/h],         is referred to here and hereinafter as nameplate load.

The maximum possible flow rate of total formaldehyde used (2) may vary according to the MDA product type to be prepared. This is generally unimportant for the present invention since the product after a change in load is essentially the same as before the change in load. If slight changes in the product composition, contrary to expectation, should lead to significant deviation in the nameplate load in the end state from that in the starting state, it is the nameplate load in the starting state m₂(N, A) that is crucial for the purposes of the present invention (i.e. more particularly for the purpose of quantification of the change in load; for details see below). The value of m₂(N, A) for a given production plant is either known from the outset (for example because the plant was designed for the production of a particular amount of a particular product type per hour) or can be easily determined by the person skilled in the art from the known boundary conditions of the plant (number and size of apparatuses present and the like) for the known operating parameters of the crucial starting state (n(1)/n(2)(A), n(7)/n(1)(A)). In practice, rather than m₂(N), the corresponding product flow rate of MDA (e.g. in t/h) which is to be expected under the assumption of the design parameters for the production plant (including n(1)/n(2) ratio, n(7)/n(1) ratio etc.) is frequently reported. Rather than nameplate load, reference is therefore also frequently made to nameplate production capacity or else nameplate capacity for short. In this connection, the terms “load” and “(production) capacity” ultimately mean the same thing, but have different reference points (flow rate of starting material in one case and flow rate of target product to be expected from this starting material in the other case).

The actual load (or—see the above elucidations—the actual production capacity or capacity for short) in the starting state and end state, defined by the actual mass flow rate of total formaldehyde used in the starting state or end state, can then be expressed in relation to the nameplate load as follows:

m ₂(A)=X(A)·m ₂(N) [e.g. in t/h];

m ₂(E)=X(E)·m ₂(N) [e.g. in t/h].

In both cases, X is a dimensionless multiplier greater than 0 and less than or equal to 1 that expresses the actual load in relation to the nameplate load. In the event of a change from production at half the nameplate load (“half-load”) to production with nameplate load (“full load”), accordingly, X(A)=½ and X(E)=1. In the case of increases in load, X(E)>X(A); in the case of reductions in load, X(E)<X(A).

A quantitative measure for the size of a change in load is the magnitude of the differential of X(E) and X(A):

|X(E)−X(A)|(=|X(A)−X(E)|)=ΔX

In all abovementioned states of operation, n(1)/n(2) is set to a value of not less than 1.6. With regard to the expression “total acidic catalyst used”, the statements made above for aniline and formaldehyde are correspondingly applicable: what is meant is the total amount of acidic catalyst. This can be (and is preferably) fed entirely to the reactor 3000-1; alternatively, it is possible to supply this reactor with just the majority of the total acidic catalyst used and to meter smaller proportions additionally into the reactors that follow in flow direction.

According to the invention, moreover, the temperature in each of the reactors downstream in flow direction (3000-2, 3000-3, . . . 3000-i) is set to a value more than 2.0° C. above T₃₀₀₀₋₁. This is also true in all states of operation, i.e. in the starting state, transition state and end state.

In any case, the aim is to largely avoid unwanted changes in the product composition resulting from the change in load.

According to the invention, the temperature of each reactor j of the reactor cascade (3000) in the end state is adjusted to a value that equates to the respective temperature in the starting state within a range of variation of ±2.0° C.; in other words, in each case:

(T _(3000-j)(A)−2.0° C.)≤T _(3000-j)(E)≤(T _(3000-j)(A)+2.0° C.).

According to the invention, moreover, the temperature in the first reactor in flow direction (3000-1) is always chosen such that it is within the range from 25.0° C. to 65.0° C., preferably 30.0° C. to 60.0° C. This is true in all states of operation, i.e. in the starting state, transition state and end state.

Moreover, during the transition state T:

(i) the temperature in the first reactor in flow direction (3000-1) from step (A-II) is set to a value that differs from the temperature in this reactor during the starting state A by a maximum of 10.0° C., preferably by a maximum of 5.0° C., more preferably by a maximum of 2.0° C., the particularly preferred maximum deviation by ±2.0° C. being regarded as the same temperature for all practical purposes (in the context of the present invention, slight nominal deviations from a temperature in the region of up to ±2.0° C. are considered to be—essentially—the same temperature; changes of greater than 2.0° C., by contrast, are considered to be significant temperature changes);

(ii) the temperature in all of the reactors downstream in flow direction (3000-2, . . . , 3000-i), by comparison with the starting state A, in each case is kept the same within a range of variation of ±2.0° C.;

(iii-1) in the case that m₂(E)>m₂(A), n(1)/n(2)(T) and n(7)/n(1)(T) are adjusted such that, at the end of the transition state, i.e. when m₂ has reached the target for m₂(E):

0.80·n(1)/n(2)(A)≤n(1)/n(2)(T)≤0.99·n(1)/n(2)(A), preferably

0.90·n(1)/n(2)(A)≤n(1)/n(2)(T)≤0.97·n(1)/n(2)(A);

and

1.01·n(7)/n(1)(A)≤n(7)/n(1)(T)≤2.00·n(7)/n(1)(A), preferably

1.05·n(7)/n(1)(A)≤n(7)/n(1)(T)≤1.50·n(7)/n(1)(A);

where it is always the case that, throughout the transition state, prior to attainment of the time at which m₂ has reached the target value for m₂(E):

0.80·n(1)/n(2)(A)≤n(1)/n(2)(T)≤1.20·n(1)/n(2)(A), preferably

0.90·n(1)/n(2)(A)≤n(1)/n(2)(T)≤1.10·n(1)/n(2)(A);

and

0.83·n(7)/n(1)(A)≤n(7)/n(1)(T)≤2.00·n(7)/n(1)(A), preferably

0.91·n(7)/n(1)(A)≤n(7)/n(1)(T)≤1.50·n(7)/n(1)(A);

(iii-2) in the case that m₂(E)<m₂(A), n(1)/n(2)(T) and n(7)/n(1)(T) are adjusted such that, at the end of the transition state, i.e. when m₂ has reached the target value for m₂(E):

1.01·n(1)/n(2)(A)≤n(1)/n(2)(T)≤1.50·n(1)/n(2)(A), preferably

1.02·n(1)/n(2)(A)≤n(1)/n(2)(T)≤1.20·n(1)/n(2)(A);

and

0.50·n(7)/n(1)(A)≤n(7)/n(1)(T)≤0.99·n(7)/n(1)(A), preferably

0.60·n(7)/n(1)(A)≤n(7)/n(1)(T)≤0.80·n(7)/n(1)(A);

where it is always the case that, throughout the transition state, prior to attainment of the time at which m₂ has reached the target value for m₂(E):

0.90·n(1)/n(2)(A)≤n(1)/n(2)(T)<1.50·n(1)/n(2)(A), preferably

0.95·n(1)/n(2)(A)≤n(1)/n(2)(T)≤1.20·n(1)/n(2)(A);

and

0.50·n(7)/n(1)(A)≤n(7)/n(1)(T)≤1.11·n(7)/n(1)(A), preferably

0.60·n(7)/n(1)(A)≤n(7)/n(1)(T)≤1.05·n(7)/n(1)(A).

According to the invention, for the reactor 3000-1 during the transition state, a significant deviation from the temperature in the starting state is possible. In this reactor, changes in the amount of heat released as a consequence of the change in load are manifested to an enhanced degree. (Significant contributions to exothermicity are firstly the heat of neutralization which is released in the reaction of the acid with the aminic components of the reaction mixture, and the acid-catalyzed rearrangement of the aminal that sets in.) This can be taken into account if required by greater tolerances in the temperature in the transition state. However, not later than at the end of the transition state (i.e. when m₂ has reached the target value for m₂(E)), the temperature in the reactor 3000-1 is adjusted back to a value which corresponds to that of the starting state within a range of variation of ±2.0° C.

The appended drawings are intended to illustrate the procedure of the invention:

FIG. 1a-c show the ratio n(1)/n(2) and the flow rate m₂, each as a function of time t.

FIG. 2a-b show the ratio n(7)/n(1) and the flow rate m₂, each as a function of time t.

These and the other drawings are merely intended to illustrate the basic principle of the invention and do not claim to be true to scale.

It is essential to the invention that, at the end of the transition state (i.e. when the target load for the end state has been established, at t=t₂), in the case of an increase in load, the molar ratio of total aniline used to total formaldehyde used (n(1)/n(2)) is lowered compared to the starting state and the molar ratio of total acidic catalyst used to total aniline used (n(7)/n(1)) is increased compared to the starting state. In the case of a decrease in load, the procedure is reversed (increasing of n(1)/n(2) and lowering of n(7)/n(1)). The procedure of the invention is illustrated in FIG. la-c using the example of an increase in load for the ratio n(1)/n(2): In the starting state, the load m₂=m₂(A). At time t=t₁ (=commencement of the transition state), the load is increased until, at time t=t₂ (=end of the transition state), it has reached the target value for the end state m₂=m₂(E); in other words—and for all conceivable configurations of the invention and not just cases shown by way of example in FIG. 1a-c —m₂(t=t₂)=m₂(E) and n(1)/n(2)(T)(t=t₂)=n(1)/n(2)(E). The lowering of n(1)/n(2) may, as shown in FIG. 1 a, be continuous. However, this is not obligatory; for example, it is also possible, as shown in FIG. 1 b, first to increase n(1)/n(2) (by increasing m₁ to a greater degree at first than m₂) and only to lower it to the target end value with a time delay, which leads to an intermediate rise in n(1)/n(2). All that is essential is that the target end value for the ratio n(1)/n(2) is established during the transition state in such a way that this target end value n(1)/n(2)=n(1)/n(2)(E) exists at time t=t₂ (i.e. at the time at which m₂ has reached the target value for m₂(E)). This is not restricted to the case of an increase in load, as shown by way of example in FIG. 1a -c, but is also true of all embodiments of the invention. The target end value n(1)/n(2)(E) can be established stepwise or, as shown in FIG. 1a -c, continuously, preference being given to continuous establishment. The same is true of the ratio n(7)/n(1) (FIG. 2a-c ). It is therefore likewise the case—again for all conceivable configurations of the invention and not just in the cases shown by way of example in FIG. 2a-c —that n(7)/n(1)(T)(t=t₂)=n(7)/n(1)(E). The respectively required adjustments in the operating parameter m₂ and at least one operating parameter selected from m₁ and m₇ (preferably, both operating parameters m₁ and m₇ are altered) are preferably commenced at the same time and more preferably conducted in such a way that the target values in each case for the end of the change in load (i.e. the time at which m₂ has reached the target value for m₂(E)) are established via continuous adjustment of the respective operating parameter.

However, a slight time delay (generally in the order of magnitude of seconds, preferably not more than 15 seconds) is possible in the required adjustments of m₁, m₂ and m₇. This is especially effected in such a way that, in the case of an increase in load, it commences first with the adjustment of m₁ (if such an adjustment, generally an increase, is required) and only then is m₂ increased with a slight time delay, in such a way that the requirements of the invention on the ratio n(1)/n(2) are complied with. In the case of a decrease in load, the reverse procedure is followed (commencement with the adjustment (i.e. in this case lowering) of m₂, and optionally adjustment, generally lowering, of m, again of course with the proviso that the requirements of the invention on the ratio n(1)/n(2) are complied with). In this embodiment too, the adjustment of m₇ required in each case is also undertaken preferably at the same time as the adjustment of m₁. If, in this embodiment, an increase in load is commenced with an increase in m₁ before the increase in m₂ is commenced, there is already a rise in the n(1)/n(2) ratio shortly before t=t₁ (i.e. shortly before commencement of the transition state as defined in accordance with the invention). It is generally the case that, when commencing with an adjustment of m₁ before commencement of the desired change (increase in load or decrease in load) of m₂ (such a time is shown in FIG. 1c and is identified there by to), the ratios n(1)/n(2)(A) and n(7)/n(1)(A), for the purposes of the present invention, are calculated on the basis of the value of m₁ before commencement of the adjustment of m₁ (cf. FIG. 1c ).

Embodiments of the invention are described in detail hereinafter. These embodiments, unless stated otherwise in the specific case, are applicable to all process regimes of the invention. It is possible here to combine various embodiments with one another as desired, unless the opposite is apparent to the person skilled in the art from the context.

FIG. 3 shows a production plant (10 000) suitable for the performance of the process of the invention.

FIG. 4 shows a detail of the reactor cascade (3000). The merely optional addition of further acidic catalyst (7) to the reactors (3000-2) to (3000-i) downstream of the reactor (3000-1) is shown by dotted arrows.

Step (A-I) of the process of the invention, provided that the other requirements of the invention are complied with, can be conducted as known in principle from the prior art. Aniline and aqueous formaldehyde solution are preferably condensed here at molar ratios in the range from 1.6 to 20, preferably 1.6 to 10 and more preferably 1.6 to 6.0, even more preferably of 1.7 to 5.5 and very exceptionally preferably of 1.8 to 5.0, at temperatures of 20.0° C. to 120.0° C., preferably 40.0° C. to 110.0° C. and more preferably 60.0° C. to 100.0° C., to give aminal and water. The reactor of step (A-I) is operated at standard pressure or under elevated pressure. There is preferably a pressure of 1.05 bar to 5.00 bar (absolute), very preferably of 1.10 bar to 3.00 bar (absolute) and most preferably of 1.20 bar to 2.00 bar (absolute). The pressure is maintained by pressure-regulating valves, or by connecting the offgas systems of aminal reactor (1000) and the overflow from the aminal separator (2000) used for phase separation on completion of reaction. The aminal separator and the outlet for the aqueous phase are preferably heated in order to prevent caking.

Suitable aniline qualities are described, for example, in EP 1 257 522 B1, EP 2 103 595 A1 and EP 1 813 598 B1. Preference is given to using technical grade qualities of formalin (aqueous solution of formaldehyde) with 30.0% by mass to 50.0% by mass of formaldehyde in water. However, formaldehyde solutions with lower or higher concentrations or else the use of gaseous formaldehyde are also conceivable.

The phase separation of organic aminal phase and aqueous phase is preferably effected at temperatures of 20.0° C. to 120.0° C., more preferably of 40.0° C. to 110.0° C. and most preferably of 60.0° C. to 100.0° C., in each case preferably at ambient pressure or at slightly elevated pressure relative to ambient pressure (elevated by up to 0.10 bar) be carried out.

Step (A-II) of the process of the invention, provided that the other requirements of the invention are complied with, can be conducted as known in principle from the prior art. The aminal is rearranged in the presence of an acidic catalyst, typically a strong mineral acid such as hydrochloric acid. Preference is given to the use of mineral acid in a molar ratio of mineral acid to aniline of 0.0010 to 0.90, preferably 0.050 to 0.50. It is of course also possible to use solid acidic catalysts as described in the literature. Formaldehyde can be added here to a mixture of aniline and acidic catalyst, and the reaction solution can be fully reacted by stepwise heating. Alternatively, aniline and formaldehyde can first be pre-reacted and then, with or without prior removal of water, admixed with the acidic catalyst or a mixture of further aniline and acidic catalyst, and then the reaction solution is fully reacted by stepwise heating. This reaction can be executed continuously or batch-wise by one of the numerous methods described in the literature (for example in EP 1 616 890 A1 or EP 127 0544 A1).

It is possible to supply the first reactor 3000-1 with further aniline and/or further formaldehyde. It is likewise possible to supply the downstream reactors 3000-2, . . . , 3000-i with small amounts of aniline and/or formaldehyde. These may each be fresh feedstocks or recycle streams from other reactors. However, the majority of total aniline used and of total formaldehyde used is introduced into the “aminal reactor” 1000. Regardless of how aniline (1) and formaldehyde (2) are distributed, optionally over various reactors, the process of the invention is preferably conducted in such a way that the molar ratio of total aniline used (1) to total formaldehyde used (2), n(1)/n(2), in all states of operation (A, T, E) always has a value of 1.6 to 2.0, preferably of 1.6 to 10, more preferably of 1.6 to 6.0, even more preferably of 1.7 to 5.5 and very exceptionally preferably of 1.8 to 5.0.

Preferably, all the acidic catalyst used (7) is fed completely to reactor 3000-1. Alternatively, it is possible to feed a portion of the total acidic catalyst used (7) to one or more of the reactors 3000-2, . . . , 3000-i downstream in flow direction.

The acidic catalyst (7) used in the process of the invention is preferably a mineral acid, especially hydrochloric acid. Suitable hydrochloric acid qualities are described, for example, in EP 1 652 835 A1.

Suitable reactors 3000-1, 3000-2, . . . 3000-i in step (A) in both process regimes are apparatuses known to those skilled in the art, such as stirred tanks and tubular reactors:

In the case of stirred tank reactors, the temperature is generally the same throughout the reactor contents, and so, for the purposes of the present invention, it does not matter where the temperature is measured. If, contrary to expectation, significant temperature differences exist through the reactor contents, the temperature measured at the exit of the reaction mixture from the reactor is the crucial temperature for the purposes of the present invention.

If there is a significant temperature gradient between entry of the reaction mixture into the reactor and exit of the reaction mixture from the reactor, as may be the case in tubular reactors, the temperature measured at the exit of the reaction mixture from the reactor is the crucial temperature for the purposes of the present invention.

Preferably, the temperature in the reactors of the reactor cascade 3000 increases from reactor 3000-1 to reactor 3000-i, i.e. the temperature in the last reactor 3000-i is preferably higher than in the first reactor 3000-1, and the temperature in two successive reactors between 3000-1 and 3000-i may also be the same, but the temperature of each reactor 3000-2, 3000-3, . . . , 3000-i is not lower than that of the preceding reactor. Preferably, however, the temperature in the reactors of the reactor cascade 3000 increases successively from reactor 3000-1 to reactor 3000-i, i.e. the temperature of each reactor 3000-2, 3000-3, . . . , 3000-i is higher than that of the preceding reactor.

It is likewise preferable to set

-   -   T₃₀₀₀₋₁ always to a value of 25.0° C. to 65.0° C., more         preferably 30.0° C. to 60.0° C., and     -   the temperature in each of the reactors downstream in flow         direction (3000-2, . . . , 3000-i) always to a value of 35.0° C.         to 200.0° C., more preferably of 50.0° C. to 180.0° C.

This makes it possible, in reactor 3000-1, to achieve a practicable compromise between the aim of minimum by-product formation, especially with regard to N-methyl- and N-formyl-MDA, on the one hand (promoted by lower temperature) and the aim of ensuring a viscosity of the reaction mixture that assures practicable processibility of the product on the other hand (promoted by higher temperature).

All the aforementioned figures for preferred temperatures are applicable in all states of operation (A, T, E).

The respective values of n(1)/n(2) and n(7)/n(1) at the end of the transition state (i.e. when m₂ has reached the target value for the end state m₂(E)) are preferably retained for the duration of production with the formaldehyde mass flow rate m₂(E).

Step (B-I), provided that the other requirements of the invention are complied with, can be conducted as known in principle from the prior art. The reaction mixture comprising the di- and polyamines of the diphenylmethane series is optionally neutralized with addition of water and/or aniline. According to the prior art, the neutralization is typically effected at temperatures of, for example, 90.0° C. to 120.0° C. without addition of further substances. It can alternatively be effected at a different temperature level in order, for example, to accelerate the degradation of troublesome by-products. Suitable bases are, for example, the hydroxides of the alkali metal and alkaline earth metal elements. Preference is given to employing aqueous NaOH. The base used for neutralization is preferably used in amounts of greater than 100%, more preferably 105% to 120%, of the amount stoichiometrically required for the neutralization of the acidic catalyst used (see EP 1 652 835 A1).

Step (B-II), provided that the other requirements of the invention are complied with, can be conducted as known in principle from the prior art. The neutralized reaction mixture comprising the di- and polyamines of the diphenylmethane series is separated into an organic phase comprising di- and polyamines of the diphenylmethane series and an aqueous phase. This can be assisted by the addition of aniline and/or water. If the phase separation is assisted by addition of aniline and/or water, they are preferably added already with vigorous mixing in the neutralization. The mixing can be effected here in mixing zones with static mixers, in stirred tanks or stirred tank cascades, or else in a combination of mixing zones and stirred tanks. The neutralized reaction mixture diluted by addition of aniline and/or water is then preferably supplied to an apparatus which, owing to its configuration and/or internals, is particularly suitable for separation into an organic phase comprising MDA and an aqueous phase, preferably phase separation or extraction apparatuses according to the prior art, as described, for example, in Mass-Transfer Operations, 3rd Edition, 1980, McGraw-Hill Book Co, p. 477 to 541, or Ullmann's Encyclopedia of Industrial Chemistry (Vol. 21, Liquid-Liquid Extraction, E. Müller et al., pages 272-274, 2012 Wiley-VCH Verlag GmbH & Co. KGaA, Weinheim, DOI: 10.1002/14356007.b03_06.pub2) or in Kirk-Othmer Encyclopedia of Chemical Technology (see “http://onlinelibrary.wiley.com/book/10.1002/0471238961”, Published Online: 15 Jun. 2007, pages 22-23) (mixer-settler cascades or settling vessels).

If, in step (A-II), during the transition state, the flow rate of acidic catalyst is increased significantly, the flow rate of base used in step (B-I) will of course also be increased correspondingly (and within the same timeframe), in order that the requirement of the invention for the stoichiometric excess of base is always fulfilled, i.e. even in the transition state. If, in step (A-II), during the transition state, the flow rate of acidic catalyst is lowered significantly, the flow rate of base used in step (B-I) will preferably also be lowered correspondingly (and likewise within the same timeframe), in order to avoid unnecessary salt burdens.

Step (B-II) is preferably followed by further workup steps, namely:

Step (B-III): washing the organic phase (11) in a washing unit (6000, “washing vessel”) with washing liquid (13), followed by

Step (B-IV): separating the mixture (14) obtained in step (B-III) in a separation unit (7000, “washing water separator”) into an organic phase (16) comprising di- and polyamines of the diphenylmethane series and an aqueous phase (15);

Step (B-V): distilling the organic phase (16) from step (B-IV) in a distillation apparatus (8000) to obtain the di- and polyamines of the diphenylmethane series (18), with removal of a stream (17) comprising water and aniline.

These steps can be conducted as known in principle from the prior art. In particular, it is preferable to conduct these steps continuously (as defined above) as well.

It is particularly preferable that there is a subsequent washing (B-III) of the organic phase (11) with water (13) and a new separation of the water phase (15) for removal of residual salt contents (preferably as described in DE-A-2549890, page 3). After exiting from the phase separation in step (B-IV), the organic phase (16) comprising di- and polyamines of the diphenylmethane series typically has a temperature of 80.0° C. to 150.0° C.

Water and aniline are separated by distillation from the organic phase thus obtained, comprising di- and polyamines of the diphenylmethane series, as described in EP 1 813 597 B1. The organic phase preferably has a composition, based on the mass of the mixture, of 5.0% to 15% by mass of water and, according to the use ratios of aniline and formaldehyde, 5.0% to 90% by mass, preferably 5.0% to 40% by mass, of aniline and 5.0% to 90% by mass, preferably 50% to 90% by mass, of di- and polyamines of the diphenylmethane series.

More preferably, the process of the invention also includes a further step (C) in which recycling of stream (17) comprising water and aniline, optionally after workup, into step (A-I) and/or, if the optional addition of further aniline (1) in step (A-II) is conducted, into step (A-II) is undertaken.

In all embodiments of the invention, the setting of the operating parameters to the target values for the end state is concluded with the end of the change in load (i.e. as soon as the mass flow rate of total formaldehyde used m₂ has reached the target value for the end state m₂(E), time t₂). The values of n(1)/n(2) and n(7)/n(1) are therefore set stepwise or continuously, preferably continuously, to the target values for the end state. In the case of stepwise setting of the respective parameter, the setting preferably comprises multiple stages, meaning that, a target end value is established via one or more intermediate values, each of which is maintained for a particular period of time, between the starting value and the target end value. The term “continuous setting” also encompasses the case that the respective parameter is first varied in the “opposite direction”, as shown in FIG. 1b -c.

In all embodiments of the invention, it is further preferable to restrict the change in load to a time not exceeding 120 minutes, i.e. to limit the period within which the mass flow rate of total formaldehyde used m₂ is adjusted proceeding from m₂(A) to the target value for the end state m₂(E) (=duration of the transition state=period of time from t₁ to t₂) is preferably to 120 minutes. The minimum duration of the transition state is preferably 1.00 minute, more preferably 5.00 minutes and most preferably 30.00 minutes.

Detailed configurations of the invention are elucidated in detail in the appended examples.

The di- and polyamines of the diphenylmethane series that are obtained in accordance with the invention can be reacted by the known methods, under inert conditions, with phosgene in an organic solvent to give the corresponding di- and polyisocyanates of the diphenylmethane series, the MDI. The phosgenation can be conducted here by any of the methods known from the prior art (e.g. DE-A-844896 or DE-A-19817691).

The procedure of the invention gives rise to the following advantages for the preparation of MDA:

-   -   i) The productivity of the MDA plant is higher because the         occurrence of off-spec material is minimized.     -   ii) There is lower formation of by-products with acridine and         acridane structure and/or N-formyl- and N-methyl-MDA. Higher         proportions of these by-products can lead to quality problems in         the MDI conversion product.

Thus, the procedure of the invention, during a non-steady state (during the transition state), enables a technically seamless change in load without subsequent outage periods in the steady state (the end state) that follows with constantly high quality of the desired MDA end product. The process of the invention also enables a rapid change in load and hence rapid reaction to events such as raw material shortage, etc.

EXAMPLES

The results outlined in the examples for the bicyclic content, the isomer composition and the content of N-methyl-4,4′-MDA are based on calculations. The calculations are based partly on theoretical models and partly on process data collected in real operational experiments, the statistical evaluation of which created a mathematical correlation of running parameters and result (e.g. bicyclic content). The content of N-formyl-4,4′-MDA is reported on the basis of operational experience values. All percentages and ppm values reported are proportions by mass based on the total mass of the respective stream of matter. The proportions by mass in the real operational experiments that gave the basis for the theoretical model were ascertained by HPLC.

Reactor temperatures are based on the temperature of the respective process product at the exit from the reactor.

The MDA prepared, in all examples, has a residual aniline content in the range from 50 ppm to 100 ppm and a water content in the range from 200 ppm to 300 ppm.

A. Reduction in Load Proceeding from Production with Nameplate Load I. Starting State: Description of the Conditions Chosen for the preparation of MDA at Nameplate Load

In a continuous reaction process, 23.20 t/h of feed aniline (containing 90.0% by mass of aniline, 1) and 9.60 t/h of 32% aqueous formaldehyde solution (corresponding to a molar ratio of aniline (1):formaldehyde (2) of 2.25:1) are mixed and converted to the aminal (3) at a temperature of 90.0° C. and a pressure of 1.40 bar (absolute) in a stirred reaction tank (1000). The reaction tank is provided with a cooler having a cooling circuit pump. The reaction mixture leaving the reaction tank is guided into a phase separation apparatus (aminal separator, 2000) (step (A-I)).

After the phase separation to remove the aqueous phase (6), the organic phase (5) is admixed in a mixing nozzle with 30% aqueous hydrochloric acid (7) (protonation level 10%, i.e. 0.10 mol of HCl is added per mole of amino groups) and run into the first rearrangement reactor (3000-1). The first rearrangement reactor (called “vacuum tank”) is operated at 50.0° C., which is ensured by means of evaporative cooling in a reflux condenser at a pressure of 104 mbar (absolute). The reflux condenser is charged with 0.50 t/h of fresh aniline. The rearrangement reaction is conducted to completion in a reactor cascade composed of a total of seven reactors at 50.0° C. to 156.0° C. (i.e. 50.0° C. in reactor 3000-1/60.0° C. in reactor 3000-2/83.0° C. in reactor 3000-3/104.0° C. in reactor 3000-4/119.0° C. in reactor 3000-5/148.0° C. in reactor 3000-6/156.0° C. in reactor 3000-7) (step (A-II)).

On completion of reaction, the reaction mixture (8-i) obtained is admixed with 32% sodium hydroxide solution in a molar ratio of 1.10:1 sodium hydroxide to HCl and reacted in a stirred neutralization vessel (4000) (step (B-I)). The temperature here is 115.0° C. The absolute pressure is 1.40 bar. The neutralized reaction mixture (10) is then separated in a neutralization separator (5000) into an aqueous lower phase (12), which is guided to a wastewater collection vessel, and into an organic phase (11) (step (B-II)).

The organic upper phase (11) is guided to the washing and washed with condensate (13) in a stirred washing vessel (6000) (step (B-III)). After the washing water (15) has been separated from the biphasic mixture (14) obtained in the washing vessel (6000) in a washing water separator (7000, step (B-IV)), the crude MDA (16) thus obtained is freed of water and aniline (removed together as stream 17) by distillation, and 17.00 t/h of MDA (18) were obtained as bottom product (step (B-V)).

MDA prepared in this way has an average composition of 45.2% 4,4′-MDA, 5.5% 2,4′-MDA, 0.3% 2,2′-MDA, i.e. a total bicyclic content of 51.0% and a proportion by mass of 2,4′-MDA of 10.8%, and also 0.3% N-methyl-4,4′-MDA and 0.3% N-formyl-4,4′-MDA, the remainder to 100% consisting essentially of higher homologs (PMDA) and isomers thereof.

II. Target End State: Production at Half-Load Example 1 (Comparative Example): Reduction in Load of the MDA Plant from Nameplate Load to Half-Load (=50% of Nameplate Load), Where the n(1)/n(2) Ratio and the n(7)/n(1) Ratio are Kept the Same, the Aniline, Formalin, Hydrochloric Acid and Sodium Hydroxide Solution Feedstocks are Reduced Simultaneously and the Temperatures in the Vacuum Tank (3000-1), the Downstream Reactors and the Exit Temperature of the Crude MDA Solution in the Last Rearrangement Reactor (3000-7) Remain the Same

The MDA plant, as described above under A.I, is operated at a production capacity of 17.0 t/h of MDA. Owing to lower product demand the production load is to be halved. For this purpose, at the same time, the feed rates of aniline and formalin to the aminal reactor (1000) are adjusted to the new production load within 120 minutes. The formalin rate is reduced to 4.80 t/h. The aniline feed rate is reduced to 11.35 t/h. At the same time, the flow rate of hydrochloric acid into the mixing nozzle in the feed to the first rearrangement reactor (3000-1) is halved. The first rearrangement reactor is still operated at 50.0° C., which is ensured by means of the evaporative cooling in the reflux condenser at 104 mbar (absolute). The reflux condenser is still charged with 0.50 t/h of fresh aniline. The rearrangement reaction is conducted to completion in a reactor cascade at 50.0° C. to 156.0° C. (50.0° C. in reactor 3000-1/60.0° C. in reactor 3000-2/83.0° C. in reactor 3000-3/104.0° C. in reactor 3000-4/119.0° C. in reactor 3000-5/148.0° C. in reactor 3000-6/156.0° C. in reactor 3000-7). On completion of reaction, the reaction mixture obtained, as described in the general conditions for preparation of MDA, is neutralized with sodium hydroxide solution, with reduction of the amount of sodium hydroxide solution within the same time window as formalin and aniline, and then worked up to give MDA (18).

40 hours after commencement of the change in load, the MDA in the feed to the MDA tank has a composition of 44.7% 4,4′-MDA, 5.2% 2,4′-MDA, 0.2% 2,2′-MDA, i.e. a total bicyclic content of 50.1% (determined by HPLC) and also 0.3% N-methyl-4,4′-MDA and 0.3% N-formyl-4,4′-MDA, the remainder to 100% consisting essentially of higher homologs (PMDA) and isomers thereof. The MDA has an elevated proportion of unwanted by-products with acridine and acridane structure. In this regard, see also section C further down.

Example 2 (Inventive): Reduction in Load of the MDA Plant from Nameplate Load to Half-Load, Where the n(1)/n(2) Ratio is Increased, the Aniline and Formalin Feedstocks are Reduced Simultaneously and Where the n(7)/n(1) Ratio is Reduced, Hydrochloric Acid and Sodium Hydroxide Solution are Reduced Simultaneously, the Temperatures in the Vacuum Tank (3000-1), the Downstream Reactors and the Exit Temperature of the Crude MDA Solution in the Last Rearrangement Reactor (3000-7) Remain the Same

The MDA plant, as described above under A.I, is operated at a production capacity of 17.0 t/h of MDA. Owing to lower product demand the production load is to be halved. For this purpose, at the same time, the feed rates of aniline and formalin to the aminal reactor (1000) are adjusted to the new production load within 120 minutes. The formalin rate is reduced to 4.80 t/h. The aniline feed rate is reduced to 11.90 t/h. The flow rate of hydrochloric acid into the mixing nozzle in the feed to the first rearrangement reactor is adjusted within the same period as aniline and formalin while reducing the protonation level to 7.0%. The first rearrangement reactor (3000-1) is still operated at 50.0° C., which is ensured by means of the evaporative cooling in the reflux condenser at 104 mbar (absolute). The reflux condenser is still charged with 0.50 t/h of fresh aniline. The rearrangement reaction is conducted to completion in a reactor cascade at 50.0° C. to 156° C. (50.0° C. in reactor 3000-1/60.0° C. in reactor 3000-2/83.0° C. in reactor 3000-3/104.0° C. in reactor 3000-4/119.0° C. in reactor 3000-5/148.0 ° C. in reactor 3000-6/156.0 ° C. in reactor 3000-7) (step (A-II)). On completion of reaction, the reaction mixture obtained, as described in the general conditions for preparation of MDA, is neutralized with sodium hydroxide solution, with reduction of the amount of sodium hydroxide solution within the same time window as formalin and aniline and HCl with retention of the molar ratio of 1.10:1 sodium hydroxide solution to HCl, and then worked up to give the desired MDA type, obtaining 8.5 t/h of MDA (18) as the bottom product from the distillation.

40 hours after commencement of the change in load, the MDA in the feed to the MDA tank has a composition of 45.9% 4,4′-MDA, 5.6% 2,4′-MDA, 0.3% 2,2′-MDA, i.e. a total bicyclic content of 51.8% and also 0.2% N-methyl-4,4′-MDA and 0.3% N-formyl-4,4′-MDA, the remainder to 100% consisting essentially of higher homologs (PMDA) and isomers thereof. The product differs only insignificantly from the MDA stream (18) which is obtained on average in the starting state as described in A.I. An elevated proportion of unwanted by-products with acridine and acridane structure is not formed.

Table 1 below compares the results from section A.

TABLE 1 Comparison of the examples from section A End state Half-load End state Starting state Example 1 Half-load Nameplate load (comp.) Example 2 (inv.) Aniline (90%) in reactor 1000 [t/h] 23.20 11.35 11.90 Aniline in reactor 3000-1 [t/h] 0.50 0.50 0.50 Formalin (32%) in reactor 1000 [t/h] 9.60 4.80 4.80 n(1)/n(2) 2.25 2.25 2.35 [n(1)/n(2) (T) (t = t₂)]/[n(1)/n(2) (A)] — 1.00 1.04 Protonation level [%] 10 10 7.0 n(7)/n(1) 0.10 0.10 0.070 [n(7)/n(1) (T) (t = t₂)]/[n(7)/n(1) (A)] — 1.00 0.70 Temp. gradient in reactor cascade 50.0 → 156.0 50.0 → 156.0 50.0 → 156.0 3000 [° C.] T(3000-1) [° C.] 50.0 50.0 50.0 T(3000-2) [° C.] 60.0 60.0 60.0 T(3000-3) [° C.] 83.0 83.0 83.0 T(3000-4) [° C.] 104.0 104.0 104.0 T(3000-5) [° C.] 119.0 119.0 119.0 T(3000-6) [° C.] 148.0 148.0 148.0 T(3000-7) [° C.] 156.0 156.0 156.0 Production capacity [t/h] 17.00 8.50 8.50 4,4′-MDA [%] 45.2 44.7 45.9 2,4′-MDA [%] 5.5 5.2 5.6 2,2′-MDA [%] 0.3 0.2 0.3 N-methyl-MDA [%] 0.3 0.3 0.2 N-formyl-MDA [%] 0.3 0.3 0.3 Bicyclic content [%] 51.0 50.1 51.8 Comment — Elevated Product has proportion of comparable unwanted by- isomer products with composition and a acridine and similar by-product acridane structure spectrum to the product in the starting state

B. Increase in Load Proceeding from Production at Half-Load I. Starting State: Description of the Conditions Chosen for the Preparation of MDA at Half-Load

(N.B. There are of course various options for operating a production plant 10 000 at half-load.) The conditions set in example 2 are one option; for examples 3 and 4 which follow, another option was chosen for the “half-load” starting state.)

The reaction is operated as described above under A.I for nameplate load with the following differences:

11.35 t/h of feed aniline (containing 90.0% by mass of aniline);

4.80 t/h of 32% aqueous formaldehyde solution (i.e. the molar ratio of aniline:formaldehyde is 2.25:1);

50.0° C. in reactor 3000-1/60.0° C. in reactor 3000-2/81.0° C. in reactor 3000-3/95.0° C. in reactor 3000-4/116.0° C. in reactor 3000-5/144.0° C. in reactor 3000-6/146.0° C. in reactor 3000-7;

Bottom product of 8.50 t/h of MDA (18).

MDA prepared in this way has an average composition of 46.3% 4,4′-MDA, 5.0% 2,4′-MDA, 0.2% 2,2′-MDA, i.e. a total bicyclic content of 51.5% and also 0.3% N-methyl-4,4′-MDA and 0.3% N-formyl-4,4′-MDA, the remainder to 100% consisting essentially of higher homologs (PMDA) and isomers thereof.

II. Target End State: Production at Nameplate Load Example 3 (Comparative Example): Increase in Load of the MDA Plant from Half-Load to Nameplate Load, Where the n(1)/n(2) Ratio and the n(7)/n(1) Ratio are Kept the Same, the Aniline, Formalin, Hydrochloric Acid and Sodium Hydroxide Solution Feedstocks are Increased Simultaneously and the Temperatures in the Vacuum Tank (3000-1), the Downstream Reactors and the Exit Temperature of the Crude MDA Solution in the Last Rearrangement Reactor (3000-7) Remain the Same

The MDA plant, as described above under B.I, is operated at a production capacity of 8.50 t/h of MDA. Owing to higher product demand the production load is to be doubled to nameplate load. For this purpose, at the same time, the feed rates of aniline and formalin to the aminal reactor are adjusted to the new production load within 120 minutes. The formalin rate was increased to 9.60 t/h. The aniline feed rate was increased to 23.20 t/h. The flow rate of hydrochloric acid into the mixing nozzle in the feed to the first rearrangement reactor is increased within the same period as aniline and formalin with retention of the protonation level of 10% (i.e. with retention of the n(7)/n(1) ratio). The first rearrangement reactor (3000-1) is still operated at 50.0° C., which is ensured by means of the evaporative cooling in the reflux condenser at 104 mbar (absolute). The reflux condenser is charged with 0.50 t/h of fresh aniline. The rearrangement reaction is still conducted to completion in the reactor cascade at 50.0° C. to 145° C. (50.0° C. in reactor 3000-1/60.0° C. in reactor 3000-2/82.0° C. in reactor 3000-3/96.0° C. in reactor 3000-4/117.0° C. in reactor 3000-5/142.0° C. in reactor 3000-6/145.0° C. in reactor 3000-7).

After the reaction, the reaction mixture obtained is admixed with 32% sodium hydroxide solution in a molar ratio of 1.10:1 sodium hydroxide solution to HCl and reacted in a stirred neutralization vessel, increasing the amount of sodium hydroxide solution within the same time window as formalin, aniline and HCl with retention of the molar ratios. The further workup is effected as described above under A.I. At the end of the transition state, 17.0 t/h of a bottom product are obtained.

Result: “MDA” thus prepared has not been completely rearranged and still contains partly rearranged products such as aminobenzylanilines, which lead to quality problems in the subsequent phosgenation to give MDI, for example distinctly elevated color values in the resulting MDI product. The MDA thus prepared, and also the resulting MDI, are off-spec product.

Example 4 (Inventive) Increase in Load of the MDA Plant from Half-Load to Nameplate Load, Where the n(1)/n(2) Ratio is Lowered, the Aniline and Formalin Feedstocks are Increased Simultaneously, and Where the n(7)/n(1) Ratio is Increased, Hydrochloric Acid and Sodium Hydroxide Solution are Increased Simultaneously, the Temperatures in the Vacuum Tank (3000-1), the Downstream Reactors and the Exit Temperature of the Crude MDA Solution in the Last Rearrangement Reactor (3000-7) Remain the Same

The MDA plant, as described above under B.I, is operated at a production capacity of 8.50 t/h of MDA. Owing to higher product demand the production load is to be doubled to nameplate load. For this purpose, at the same time, the feed rates of aniline and formalin to the aminal reactor are adjusted to the new production load within 120 minutes. The formalin rate is increased to 10.00 t/h. The aniline feed rate is increased to 23.20 t/h. The flow rate of hydrochloric acid into the mixing nozzle in the feed to the first rearrangement reactor is adjusted within the same period as aniline and formalin with an increase of the protonation level to 13%. The first rearrangement reactor (3000-1) is still operated at 50.0° C., which is ensured by means of the evaporative cooling in the reflux condenser at 104 mbar (absolute). The reflux condenser is charged with 0.50 t/h of fresh aniline. The rearrangement reaction is conducted to completion in the reactor cascade at 50.0° C. to 146° C. (50.0° C. in reactor 3000-1/60.0° C. in reactor 3000-2/81.0° C. in reactor 3000-3/95.0° C. in reactor 3000-4/117.0° C. in reactor 3000-5/144.0° C. in reactor 3000-6/146.0° C. in reactor 3000-7) (step A-II)).

After the reaction, the reaction mixture obtained is admixed with 32% sodium hydroxide solution in a molar ratio of 1.10:1 sodium hydroxide solution to HCl and reacted in a stirred neutralization vessel, increasing the amount of sodium hydroxide solution within the same time window as formalin, aniline and hydrochloric acid. The temperature is 115.0° C. The absolute pressure is 1.40 bar. The further workup is effected as described further up under A.I. At the end of the transition state, the bottom product obtained is 17.00 t/h of MDA (18).

20 hours after commencement of the change in load, the MDA in the feed to the MDA tank has a composition of 45.8% 4,4′-MDA, 4.9% 2,4′-MDA, 0.2% 2,2′-MDA, i.e. a total bicyclic content of 50.9% and also 0.3% N-methyl-4,4′-MDA and 0.3% N-formyl-4,4′-MDA, the remainder to 100% consisting essentially of higher homologs (PMDA) and isomers thereof. The product differs only insignificantly from the MDA stream (18) which is obtained on average in the starting state as described in B.I. An elevated proportion of unwanted by-products with acridine and acridane structure is not formed.

Table 2 below compares the results from section B.

TABLE 2 Comparison of the examples from section B End state End state Starting state Nameplate load Nameplate load Half-load Example 3 (comp.) Example 4 (inv.) Aniline (90%) in reactor 1000 [t/h] 11.35 23.20 23.20 Aniline in reactor 3000-1 [t/h] 0.50 0.50 0.50 Formalin (32%) in reactor 1000 [t/h] 4.80 9.60 10.00 n(1)/n(2) 2.25 2.25 2.16 [n(1)/n(2) (T) (t = t₂)]/[n(1)/n(2) (A)] — 1.00 0.96 Protonation level [%] 10 10 13 n(7)/n(1) 0.10 0.10 0.13 [n(7)/n(1) (T) (t = t₂)]/[n(7)/n(1) (A)] — 1.00 1.30 Temp. gradient in reactor cascade 50.0 → 146.0 50.0 → 145.0 50.0 → 146.0 3000 [° C.] pT(3000-1) [° C.] 50.0 50.0 50.0 T(3000-2) [° C.] 60.0 60.0 60.0 T(3000-3) [° C.] 81.0 82.0 81.0 T(3000-4) [° C.] 95.0 96.0 95.0 T(3000-5) [° C.] 116.0 117.0 117.0 T(3000-6) [° C.] 144.0 142.0 144.0 T(3000-7) [° C.] 146.0 145.0 146.0 Production capacity [t/h] 8.50 17.00 17.00 4,4′-MDA [%] 46.3 — 45.8 2,4′-MDA [%] 5.0 — 4.9 2,2′-MDA [%] 0.2 — 0.2 N-methyl [%] 0.3 — 0.3 N-formyl [%] 0.3 — 0.3 Bicyclic [%] 51.5 — 50.9 Comment — MDA has been Product has incompletely comparable rearranged and still isomer contains partly composition and a rearranged similar by-product products such as spectrum to the aminobenzylanilines product in the which lead to starting state quality problems in the subsequent phosgenation to give MDI, for example distinctly elevated color values

C. Fundamental Experiments for Formation of By-Products with Acridine and Acridane Structure

In a series of experiments, the amount of aqueous 30% hydrochloric acid required in each case to achieve the desired protonation level (see table 3 below) was added to a 2,2′-MDA solution in aniline preheated to 100.0° C. The 2,2′-MDA concentration in each of the individual experiments was 1.0% by mass; in addition, the solutions contained octadecane as an internal standard for gas chromatography (GC) analysis. The resulting mixture was transferred as quickly as possible by means of a peristaltic pump to a Mai glass autoclave preheated to 120.0° C. and heated to the reaction temperature envisaged (see table 3). On attainment of the desired reaction temperature, the first sample was taken (time=zero). In order to monitor the progress of the reaction, further samples were taken after 30, 60, 120 and 240 minutes and analyzed by means of GC analysis. Reaction conditions and experimental results are collated in table 3 below.

TABLE 3 Laboratory experiments for formation of the acridane and acridine secondary components Reaction Protonation Sum total temperature level Time (acridine + acridane) Experiment [° C.] [%] [min] [ppm] 1 160° C. 10% 15 79 30 208 60 421 120 925 240 1775 2 170° C. 25% 0 977 30 2554 60 3950 120 5647 240 8894 3 160° C. 25% 0 364 30 930 60 1594 120 2681 240 4354 4 180 10% 0 606 30 1630 60 2613 120 3932 240 5294 5 170 10% 0 292 30 794 60 1533 120 2406 240 3932 6 160 5% 0 0 30 0 60 0 120 0 240 0 7 170 5% 0 0 30 0 60 0 120 0 240 218 8 180 5% 0 0 30 402 60 834 120 1355 240 2380

It is found that, with rising temperature and rising hydrochloric acid concentration, the formation of the acridine and acridane secondary components from 2,2′-MDA occurs to an increased degree. 

1. A process for preparing di- and polyamines of the diphenylmethane series from aniline (1) and formaldehyde (2) in a production plant (10 000), where the molar ratio of total aniline used (1) to total formaldehyde used (2), n(1)/n(2), is always greater than 1.6, comprising: (A-I) reacting aniline (1) and formaldehyde (2) in the absence of an acidic catalyst to obtain a reaction mixture (4) comprising an aminal (3), and then at least partly separating an aqueous phase (6) from the reaction mixture (4) to obtain an organic phase (5) comprising the aminal (3); (A-II) contacting the organic phase (5) which comprises the aminal and is obtained in step (A-I) with an acidic catalyst (7) in a reactor cascade (3000) composed of i reactors connected in series (3000-1, 3000-2, . . . , 3000-i), where i is a natural number from 2 to 10, wherein the first reactor (3000-1) in flow direction is operated at a temperature T3000-1 in the range from 25.0° C. to 65.0° C., and is charged with stream (5) and acidic catalyst (7) and optionally with further aniline (1) and/or further formaldehyde (2), and every reactor downstream in flow direction (3000-2, . . . , 3000-i) is operated at a temperature of more than 2.0° C. above T3000-1 and is charged with the reaction mixture obtained in the reactor immediately upstream; (B) isolating the di- and polyamines of the diphenylmethane series from the reaction mixture (8-i) obtained from step (A-II) in the last reactor (3000-i) by a process comprising: (B-I) adding a stoichiometric excess of base (9), based on the total amount of acidic catalyst used (7), to the reaction mixture (8-i) obtained in the last reactor (3000-i) in step (A-II) to obtain a reaction mixture (10); and (B-II) separating the reaction mixture (10) obtained in step (B-I) into an organic phase (11) comprising di- and polyamines of the diphenylmethane series and an aqueous phase (12); wherein in the event of a change in the production capacity from a starting state A with a mass flow rate in the starting state of total aniline used of m₁(A)≠0, a mass flow rate in the starting state of total formaldehyde used of m₂(A)=X(A)·m₂(N), where X(A) is a dimensionless number >0 and ≤1 and m₂(N) denotes the nameplate load of the production plant (10 000), a molar ratio in the starting state of total aniline used (1) to total formaldehyde used (2) of n(1)/n(2)(A) and a molar ratio in the starting state of total acidic catalyst used to total aniline used of n(7)/n(1)(A) to an end state E with a mass flow rate in the end state of total aniline used of m₁(E)≠0, a mass flow rate in the end state of total formaldehyde used of m₂(E)=X(E)·m₂(N), where X(E) is a dimensionless number >0 and ≤1, a molar ratio in the end state of total aniline used (1) to total formaldehyde used (2) of n(1)/n(2)(E) and a molar ratio in the end state of total acidic catalyst used to total aniline used of n(7)/n(1)(E); by a quantity ΔX=|X(E)−X(A)|, with ΔX≥0.10, wherein the process comprises at least one change in production capacity that commences at a time t₁ and has concludedconcludes at a time t₂, where the following is true of the temperature of each reactor j of the reactor cascade (3000), T_(3000-j) is: (T _(3000-j)(A)−2.0° C.)≤T _(3000-j)(E)≤(T _(3000-j)(A)+2.0° C.) wherein, in the period from t₁ to t₂, the transition state T, with a molar ratio of total aniline used (1) to total formaldehyde used (2) of n(1)/n(2)(T) and a molar ratio of total acidic catalyst used to total aniline used of n(7)/n(1)(T), (i) the temperature in the first reactor (3000-1) in flow direction from step (A-II) is adjusted to a value that differs from the temperature in that reactor during the starting state A by not more than 10.0° C.; (ii) the temperature in all reactors downstream in flow direction (3000-2, . . . , 3000-i), by comparison with the starting state A, is kept the same in each case within a range of variation of ±2.0° C.; (iii-1) in the case that m₂(E)>m₂(A), n(1)/n(2)(T) and n(7)/n(1)(T) are adjusted such that, at time t₂: 0.80·n(1)/n(2)(A)≤n(1)/n(2)(T)≤0.99·n(1)/n(2)(A); and 0.83·n(7)/n(1)(A)≤n(7)/n(1)(T)≤2.00·n(7)/n(1)(A); where it is always the case throughout the transition state prior to attainment of time t₂ that: 0.80·n(1)/n(2)(A)≤n(1)/n(2)(T)≤1.20·n(1)/n(2)(A); and 0.83·n(7)/n(1)(A)≤n(7)/n(1)(T)≤2.00·n(7)/n(1)(A); (iii-2) in the case that m₂(E)<m₂(A), n(1)/n(2)(T) and n(7)/n(1)(T) are adjusted such that, at time t₂: 1.01·n(1)/n(2)(A)≤n(1)/n(2)(T)≤1.50·n(1)/n(2)(A); and 0.50·n(7)/n(1)(A)≤n(7)/n(1)(T)≤0.99·n(7)/n(1)(A); where it is always the case throughout the transition state prior to attainment of time t₂ that: 0.90·n(1)/n(2)(A)≤n(1)/n(2)(T)≤1.50·n(1)/n(2)(A); and 0.50·n(7)/n(1)(A)≤n(7)/n(1)(T)≤1.11·n(7)/n(1)(A).
 2. The process of claim 1, in which the temperature in the reactors of the reactor cascade 3000 increases from reactor 3000-1 to reactor 3000-i in all states of operation (A, T, E).
 3. The process of claim 1, in which it is always the case that T3000-1 is set to a value in the range from 25.0° C. to 65.0° C. and the temperature in each of the reactors downstream in flow direction (3000-2, . . . , 3000-i) is set to a value in the range from 35.0° C. to 200.0° C.
 4. The process of claim 3, in which it is always the case that T3000-1 is set to a value in the range from 30.0° C. to 60.0° C. and the temperature in each of the reactors downstream in flow direction (3000-2, . . . , 3000-i) is set to a value in the range from 50.0° C. to 180.0° C.
 5. The process of claim 1, in which the acidic catalyst (7) is a mineral acid.
 6. The process of claim 1, in which step (B) further comprises: (B-III) washing the organic phase (11) with washing liquid (13); (B-IV) separating the mixture (14) obtained in step (B-III) into an organic phase (16) comprising di- and polyamines of the diphenylmethane series and an aqueous phase (15); and (B-V) distilling the organic phase (16) from step (B-IV) to obtain the di- and polyamines of the diphenylmethane series (18), with removal of a stream (17) comprising water and aniline.
 7. The process of claim 6, additionally comprising: (C) recycling stream (17), optionally after workup, into step (A-I) and/or, if the optional addition of further aniline (1) in step (A-II) is conducted, into step (A-II).
 8. The process of claim 1, in which the molar ratio of total aniline used (1) to total formaldehyde used (2), n(1)/n(2), in all states of operation (A, T, E) is adjusted to a value of 1.6 to
 20. 9. The process of claim 1, in which the target values of n(1)/n(2) and n(7)/n(1) for the end state E are established continuously in the transition state.
 10. The process of claim 1, in which the values for n(1)/n(2) and n(7)/n(1) that exist in each case at time t₂ are retained for the duration of the production with the formaldehyde mass flow rate m₂(E).
 11. The process of claim 1, in which the period from t₁ to t₂ lasts from 1.00 minute to 120 minutes. 